Flexible operation of oligomerization process

ABSTRACT

A process and apparatus for controlling the amount of gasoline and distillate obtained from an oligomerate product is described. The process includes oligomerizing olefins in a feed stream to produce a mixture of an oligomerate product and unreacted feed. The mixture is separated into a light stream comprising the unreacted feed and the oligomerate product. Some or all of the oligomerate product is hydrogenated to produce hydrotreated oligomerate product. The hydrotreated oligomerate product can be separated into gasoline and distillate. A recycle stream comprising at least one of non-hydrotreated oligomerate product and hydrotreated oligomerate product is introduced into the oligomerization zone. The flow rates of non-hydrotreated oligomerate product and hydrotreated oligomerate product in the recycle stream are controlled to control the amount of gasoline and distillate product produced.

BACKGROUND

When oligomerizing light olefins within a refinery, there is frequently a desire to have the flexibility to make high octane gasoline, high cetane diesel, or combination of both. However, catalysts that make high octane gasoline typically make product that is highly branched and within the gasoline boiling point range. This product is very undesirable for diesel. In addition, catalysts that make high cetane diesel typically make product that is more linear and in the distillate boiling point range. This results in less and poorer quality gasoline due to the more linear nature of the product which has a lower octane value.

When oligomerizing olefins from a fluid catalytic cracking (FCC) unit, there is often the desire to maintain a liquid phase within the oligomerization reactors. A liquid phase helps with catalyst stability by acting as a solvent to wash the catalyst of heavier species produced. In addition, the liquid phase provides a higher concentration of olefins to the catalyst surface to achieve a higher catalyst activity. Typically, this liquid phase in the reactor is maintained by hydrogenating some of the heavy olefinic product and recycling this paraffinic product to the reactor inlet.

For example the C5+ olefins can react further to form distillate range material. The reaction is such that the light olefins (e.g., C3 to C6) are the most reactive followed by the C8 olefins. The heavier olefins (C9+) are less reactive and do not tend to form heavier compounds. Thus, by recycling the C5+ olefin stream, the C6-8 olefins are further reacted and will form distillate range material. Typically, this liquid phase in the reactor is maintained by hydrogenating some of the heavy olefinic product and recycling this paraffinic product to the reactor inlet.

The net C5+ stream is often sent to a downstream hydrotreater where the olefins are saturated. This is often a more desirable product than the olefin stream. The cetane of the saturated distillate product is higher than the olefin product and is more valuable. The refiner is also limited in how much of the olefins can be blended in the gasoline.

There is a need for a flexible process that allows the refiner to switch easily from distillate to gasoline as needed to meet market demands.

SUMMARY OF THE INVENTION

One aspect of the invention is a process for controlling an amount of gasoline and distillate obtained from an oligomerate product. In one embodiment, the process includes introducing an oligomerization feed stream comprising at least one of C4 or C5 olefins into an oligomerization zone under oligomerization conditions in the presence of an oligomerization catalyst to oligomerize olefins in the oligomerization feed stream in liquid phase to produce a mixture of an oligomerate product comprising C5+ hydrocarbons and unreacted C4 or C5 feed olefins. The mixture is separated into a light stream comprising the unreacted C4 or C5 feed olefins and a stream of the oligomerate product in a first separation zone. The stream of oligomerate product is optionally divided into a first portion and a second portion, the first portion being a non-hydrotreated oligomerate product. At least the second portion of the stream of the oligomerate product is hydrogenated to produce a stream of hydrotreated oligomerate product. A recycle stream is introduced into the oligomerization zone, the recycle stream comprising at least one of the optional first portion of the non-hydrotreated oligomerate product and a portion of the hydrotreated oligomerate product. The flow rate of the optional first portion of the non-hydrotreated oligomerate product and the flow rate of the recycle portion of the hydrotreated oligomerate product in the recycle stream are controlled to control the amount of gasoline and distillate product produced.

Another aspect of the invention is an apparatus for controlling the production of diesel and gasoline. In one embodiment, the apparatus includes an oligomerization reaction zone with an inlet and an outlet; a first separation zone with an inlet, an overhead outlet, and a bottoms outlet; the inlet of the separation zone being in fluid communication with the outlet of the oligomerization zone; a hydrogenation zone having an inlet and an outlet, the inlet of the hydrogenation zone being in fluid communication with the bottoms outlet of the separation zone; a second separation zone with an inlet, an overhead outlet, and a bottoms outlet, the inlet of the second separation zone being in fluid communication with the outlet of the hydrogenation zone; the bottoms outlet of the first separation zone being in fluid communication with the inlet of the oligomerization zone and the outlet of the hydrogenation zone being in fluid communication with the inlet of the oligomerization zone; and a controller for controlling a flow from the bottoms outlet of the first separation zone to the inlet of the oligomerization zone and a flow from the outlet of the hydrogenation zone to the inlet of the oligomerization zone.

BRIEF DESCRIPTION OF THE DRAWING

The FIGURE is a schematic drawing of one embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention meets that need by providing a process that can be shifted from producing a maximum amount of distillate to a maximum amount of gasoline and anywhere between the two easily using existing equipment.

The process involves integrating the hydrogenation zone into the oligomerization process. The total amount of recycle to the oligomerization zone remains constant, but the ratio of oligomerization product to the hydrotreated oligomerization product is varied. When increased amounts of gasoline are desired, the amount of hydrotreated oligomerate product is increased. If more distillate product is desired, the amount of oligomerate product is increased. The amounts can be varied from 100% oligomerate product for maximum distillate production to 100% hydrotreated oligomerate product for maximum gasoline production.

As used herein, the term “stream” can include various hydrocarbon molecules and other substances. Moreover, the term “stream comprising Cx hydrocarbons” or “stream comprising Cx olefins” can include a stream comprising hydrocarbon or olefin molecules, respectively, with “x” number of carbon atoms, suitably a stream with a majority of hydrocarbons or olefins, respectively, with “x” number of carbon atoms and preferably a stream with at least 75 wt-% hydrocarbons or olefin molecules, respectively, with “x” number of carbon atoms. Moreover, the term “stream comprising Cx+ hydrocarbons” or “stream comprising Cx+ olefins” can include a stream comprising a majority of hydrocarbon or olefin molecules, respectively, with more than or equal to “x” carbon atoms and suitably less than 10 wt-% and preferably less than 1 wt-% hydrocarbon or olefin molecules, respectively, with x−1 carbon atoms. Lastly, the term “Cx− stream” can include a stream comprising a majority of hydrocarbon or olefin molecules, respectively, with less than or equal to “x” carbon atoms and suitably less than 10 wt-% and preferably less than 1 wt-% hydrocarbon or olefin molecules, respectively, with x+1 carbon atoms.

As used herein, the term “zone” can refer to an area including one or more equipment items and/or one or more sub-zones. Equipment items can include one or more reactors or reactor vessels, heaters, exchangers, pipes, pumps, compressors, controllers and columns. Additionally, an equipment item, such as a reactor, dryer, or vessel, can further include one or more zones or sub-zones.

As used herein, the term “substantially” can mean an amount of at least generally about 70%, preferably about 80%, and optimally about 90%, by weight, of a compound or class of compounds in a stream.

As used herein, the term “gasoline” can include hydrocarbons having a boiling point temperature in the range of about 25 to about 200° C. at atmospheric pressure.

As used herein, the term “diesel” or “distillate” can include hydrocarbons having a boiling point temperature in the range of about 150 to about 400° C. and preferably about 200 to about 400° C.

As used herein, the term “vacuum gas oil” (VGO) can include hydrocarbons having a boiling temperature in the range of from 343° to 552° C.

As used herein, the term “vapor” can mean a gas or a dispersion that may include or consist of one or more hydrocarbons.

As used herein, the term “overhead stream” can mean a stream withdrawn at or near a top of a vessel, such as a column.

As used herein, the term “bottom stream” can mean a stream withdrawn at or near a bottom of a vessel, such as a column.

As depicted, process flow lines in the FIGURES can be referred to interchangeably as, e.g., lines, pipes, feeds, gases, products, discharges, parts, portions, or streams.

As used herein, “bypassing” with respect to a vessel or zone means that a stream does not pass through the zone or vessel bypassed although it may pass through a vessel or zone that is not designated as bypassed.

The term “communication” means that material flow is operatively permitted between enumerated components.

The term “downstream communication” means that at least a portion of material flowing to the subject in downstream communication may operatively flow from the object with which it communicates.

The term “upstream communication” means that at least a portion of the material flowing from the subject in upstream communication may operatively flow to the object with which it communicates.

The term “direct communication” means that flow from the upstream component enters the downstream component without undergoing a compositional change due to physical fractionation or chemical conversion.

The term “column” means a distillation column or columns for separating one or more components of different volatilities. Unless otherwise indicated, each column includes a condenser on an overhead of the column to condense and reflux a portion of an overhead stream back to the top of the column and a reboiler at a bottom of the column to vaporize and send a portion of a bottom stream back to the bottom of the column. Feeds to the columns may be preheated. The top pressure is the pressure of the overhead vapor at the outlet of the column. The bottom temperature is the liquid bottom outlet temperature. Overhead lines and bottom lines refer to the net lines from the column downstream of the reflux or reboil to the column.

As used herein, the term “boiling point temperature” means atmospheric equivalent boiling point (AEBP) as calculated from the observed boiling temperature and the distillation pressure, as calculated using the equations furnished in ASTM D1160 appendix A7 entitled “Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures”.

As used herein, “taking a stream from” means that some or all of the original stream is taken.

As used herein, “about” means within 10% of the value, or within 5%, or within 1%.

As used herein, zeolites may be referred to by proper name, such as ZSM-5, or by structure type code, such as MFI. These three letter codes indicate atomic connectivity and hence pore size, shape and connectivity for the various known zeolites. The list of these codes may be found in the Atlas of Zeolite Framework Types, which is maintained by the International Zeolite Association Structure Commission at http://www.iza-structure.org/databases/.

The products of olefin oligomerization are usually mixtures of, for example, olefin dimers, trimers, and higher oligomers. Further, each olefin oligomer is itself usually a mixture of isomers, both skeletal and in double bond location. Highly branched isomers are less reactive than linear or lightly branched materials in many of the downstream reactions for which oligomers are used as feedstocks. This is also true of isomers in which access to the double bond is sterically hindered. Olefin types of the oligomers can be denominated according to the degree of substitution of the double bond, as follows:

TABLE 1 Olefin Type Structure Description I R—HC═CH₂ Monosubstituted II R—HC═CH—R Disubstituted III RRC═CH₂ Disubstituted IV RRC═CHR Trisubstituted V RRC═CRR Tetrasubstituted

wherein R represents an alkyl group, each R being the same or different. Type I compounds are sometimes described as α- or vinyl olefins and Type III as vinylidene olefins. Type IV is sometimes subdivided to provide a Type IVA, in which access to the double bond is less hindered, and Type IVB where it is more hindered.

The present invention involves an apparatus and process that can be used to make the maximum amount of gasoline, the maximum amount of distillate, or any amount between the two. Varying amounts of gasoline and distillate are produced depending on the way the process is run and the amount of oligomerate product recycle and hydrotreated oligomerate product recycle.

The apparatus and process may be described with reference to components shown in the FIGURE: a fluid catalytic cracking (FCC) zone 20, an FCC recovery zone 100, a purification zone 110, an oligomerization zone 130, and an oligomerization recovery zone 200. Many configurations of the present invention are possible, but specific embodiments are presented herein by way of example. All other possible embodiments for carrying out the present invention are considered within the scope of the present invention.

The fluid catalytic cracking zone 20 may comprise a first FCC reactor 22, a regenerator vessel 30, and an optional second FCC reactor 70.

A conventional FCC feedstock and higher boiling hydrocarbon feedstock are a suitable FCC hydrocarbon feed 24 to the first FCC reactor. The most common of such conventional feedstocks is a VGO. Higher boiling hydrocarbon feedstocks to which this invention may be applied include heavy bottom from crude oil, heavy bitumen crude oil, shale oil, tar sand extract, deasphalted residue, products from coal liquefaction, atmospheric and vacuum reduced crudes and mixtures thereof. The FCC feed 24 may include a recycle stream 280 to be described later.

The first FCC reactor 22 may include a first reactor riser 26 and a first reactor vessel 28. A regenerator catalyst pipe 32 delivers regenerated catalyst from the regenerator vessel 30 to the reactor riser 26. A fluidization medium such as steam from a distributor 34 urges a stream of regenerated catalyst upwardly through the first reactor riser 26. At least one feed distributor injects the first hydrocarbon feed in a first hydrocarbon feed line 24, preferably with an inert atomizing gas such as steam, across the flowing stream of catalyst particles to distribute hydrocarbon feed to the first reactor riser 26. Upon contacting the hydrocarbon feed with catalyst in the first reactor riser 26 the heavier hydrocarbon feed cracks to produce lighter gaseous cracked products while coke is deposited on the catalyst particles to produce spent catalyst.

The resulting mixture of gaseous product hydrocarbons and spent catalyst continues upwardly through the first reactor riser 26 and are received in the first reactor vessel 28 in which the spent catalyst and gaseous product are separated. Disengaging arms discharge the mixture of gas and catalyst from a top of the first reactor riser 26 through outlet ports 36 into a disengaging vessel 38 that effects partial separation of gases from the catalyst. A transport conduit carries the hydrocarbon vapors, stripping media and entrained catalyst to one or more cyclones 42 in the first reactor vessel 28 which separates spent catalyst from the hydrocarbon gaseous product stream. Gas conduits deliver separated hydrocarbon cracked gaseous streams from the cyclones 42 to a collection plenum 44 for passage of a cracked product stream to a first cracked product line 46 via an outlet nozzle and eventually into the FCC recovery zone 100 for product recovery.

Diplegs discharge catalyst from the cyclones 42 into a lower bed in the first reactor vessel 28. The catalyst with adsorbed or entrained hydrocarbons may eventually pass from the lower bed into a stripping section 48 across ports defined in a wall of the disengaging vessel 38. Catalyst separated in the disengaging vessel 38 may pass directly into the stripping section 48 via a bed. A fluidizing distributor delivers inert fluidizing gas, typically steam, to the stripping section 48. The stripping section 48 contains baffles or other equipment to promote contacting between a stripping gas and the catalyst. The stripped spent catalyst leaves the stripping section 48 of the disengaging vessel 38 of the first reactor vessel 28 stripped of hydrocarbons. A first portion of the spent catalyst, preferably stripped, leaves the disengaging vessel 38 of the first reactor vessel 28 through a spent catalyst conduit 50 and passes into the regenerator vessel 30. A second portion of the spent catalyst may be recirculated in recycle conduit 52 from the disengaging vessel 38 back to a base of the first riser 26 at a rate regulated by a slide valve to recontact the feed without undergoing regeneration.

The first riser 26 can operate at any suitable temperature, and typically operates at a temperature of about 150° to about 580° C. at the riser outlet 36. The pressure of the first riser is from about 69 to about 517 kPa (gauge) (10 to 75 psig) but typically less than about 275 kPa (gauge) (40 psig). The catalyst-to-oil ratio, based on the weight of catalyst and feed hydrocarbons entering the riser, may range up to 30:1 but is typically between about 4:1 and about 10:1. Steam may be passed into the first reactor riser 26 and first reactor vessel 28 at a rate between about 2 and about 7 wt-% for maximum gasoline production and about 10 to about 15 wt % for maximum light olefin production. The average residence time of catalyst in the riser may be less than about 5 seconds.

The catalyst in the first reactor 22 can be a single catalyst or a mixture of different catalysts. Usually, the catalyst includes two catalysts, namely a first FCC catalyst, and a second FCC catalyst. Such a catalyst mixture is disclosed in, e.g., U.S. Pat. No. 7,312,370 B2. Generally, the first FCC catalyst may include any of the well-known catalysts that are used in the art of FCC. Preferably, the first FCC catalyst includes a large pore zeolite, such as a Y-type zeolite, an active alumina material, a binder material, including either silica or alumina, and an inert filler such as kaolin.

Typically, the zeolites appropriate for the first FCC catalyst have a large average pore size, usually with openings of greater than about 0.7 nm in effective diameter defined by greater than about 10, and typically about 12, member rings. Suitable large pore zeolite components may include synthetic zeolites such as X and Y zeolites, mordenite and faujasite. A portion of the first FCC catalyst, such as the zeolite portion, can have any suitable amount of a rare earth metal or rare earth metal oxide.

The second FCC catalyst may include a medium or smaller pore zeolite catalyst, such as exemplified by at least one of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48, and other similar materials. Other suitable medium or smaller pore zeolites include ferrierite, and erionite. Preferably, the second component has the medium or smaller pore zeolite dispersed on a matrix including a binder material such as silica or alumina and an inert filler material such as kaolin. These catalysts may have a crystalline zeolite content of about 10 to about 50 wt-% or more, and a matrix material content of about 50 to about 90 wt-%. Catalysts containing at least about 40 wt-% crystalline zeolite material are typical, and those with greater crystalline zeolite content may be used. Generally, medium and smaller pore zeolites are characterized by having an effective pore opening diameter of less than or equal to about 0.7 nm and rings of about 10 or fewer members. Preferably, the second FCC catalyst component is an MFI zeolite having a silicon-to-aluminum ratio greater than about 15. In one exemplary embodiment, the silicon-to-aluminum ratio can be about 15 to about 35.

The total catalyst mixture in the first reactor 10 may contain about 1 to about 25 wt-% of the second FCC catalyst, including a medium to small pore crystalline zeolite, with greater than or equal to about 7 wt-% of the second FCC catalyst being preferred. When the second FCC catalyst contains about 40 wt-% crystalline zeolite with the balance being a binder material, an inert filler, such as kaolin, and optionally an active alumina component, the catalyst mixture may contain about 0.4 to about 10 wt-% of the medium to small pore crystalline zeolite with a preferred content of at least about 2.8 wt-%. The first FCC catalyst may comprise the balance of the catalyst composition. The high concentration of the medium or smaller pore zeolite as the second FCC catalyst of the catalyst mixture can improve selectivity to light olefins. In one exemplary embodiment, the second FCC catalyst can be a ZSM-5 zeolite and the catalyst mixture can include about 0.4 to about 10 wt-% ZSM-5 zeolite excluding any other components, such as binder and/or filler.

The regenerator vessel 30 is in downstream communication with the first reactor vessel 28. In the regenerator vessel 30, coke is combusted from the portion of spent catalyst delivered to the regenerator vessel 30 by contact with an oxygen-containing gas such as air to regenerate the catalyst. The spent catalyst conduit 50 feeds spent catalyst to the regenerator vessel 30. The spent catalyst from the first reactor vessel 28 usually contains carbon in an amount of from 0.2 to 2 wt-%, which is present in the form of coke. An oxygen-containing combustion gas, typically air, enters the lower chamber 54 of the regenerator vessel 30 through a conduit and is distributed by a distributor 56. As the combustion gas enters the lower chamber 54, it contacts spent catalyst entering from spent catalyst conduit 50 and lifts the catalyst at a superficial velocity of combustion gas in the lower chamber 54 of perhaps at least 1.1 m/s (3.5 ft/s) under fast fluidized flow conditions. In an embodiment, the lower chamber 54 may have a catalyst density of from 48 to 320 kg/m³ (3 to 20 lb/ft³) and a superficial gas velocity of 1.1 to 2.2 m/s (3.5 to 7 ft/s). The oxygen in the combustion gas contacts the spent catalyst and combusts carbonaceous deposits from the catalyst to at least partially regenerate the catalyst and generate flue gas.

The mixture of catalyst and combustion gas in the lower chamber 54 ascends through a frustoconical transition section to the transport, riser section of the lower chamber 54. The mixture of catalyst particles and flue gas is discharged from an upper portion of the riser section into the upper chamber 60. Substantially completely or partially regenerated catalyst may exit the top of the transport, riser section 58. Discharge is effected through a disengaging device 58 that separates a majority of the regenerated catalyst from the flue gas. The catalyst and gas exit downwardly from the disengaging device 58. The sudden loss of momentum and downward flow reversal cause a majority of the heavier catalyst to fall to the dense catalyst bed and the lighter flue gas and a minor portion of the catalyst still entrained therein to ascend upwardly in the upper chamber 60. Cyclones 62 further separate the catalyst from the ascending gas and deposits the catalyst through dip legs into a dense catalyst bed. Flue gas exits the cyclones 62 through a gas conduit and collects in a plenum 64 for passage to an outlet nozzle of regenerator vessel 30. Catalyst densities in the dense catalyst bed are typically kept within a range of from about 640 to about 960 kg/m³ (40 to 60 lb/ft³).

The regenerator vessel 30 typically has a temperature of about 594° C. to about 704° C. (1100° F. to 1300° F.) in the lower chamber 54 and about 649° C. to about 760° C. (1200° F. to 1400° F.) in the upper chamber 60. Regenerated catalyst from dense catalyst bed is transported through regenerated catalyst pipe 32 from the regenerator vessel 30 back to the first reactor riser 26 through the control valve where it again contacts the first feed in line 24 as the FCC process continues. The first cracked product stream in the first cracked product line 46 from the first reactor 10, relatively free of catalyst particles and including the stripping fluid, exits the first reactor vessel 28 through an outlet nozzle. The first cracked products stream in the line 46 may be subjected to additional treatment to remove fine catalyst particles or to further prepare the stream prior to fractionation. The line 46 transfers the first cracked products stream to the FCC recovery zone 100, which is in downstream communication with the FCC zone 20. The FCC recovery zone 100 typically includes a main fractionation column and a gas recovery section. The FCC recovery zone can include many fractionation columns and other separation equipment. The FCC recovery zone 100 can recover a propylene product stream in propylene line 102, a gasoline stream in gasoline line 104, a light olefin stream in light olefin line 106, and an LCO stream in LCO line 107, among others, from the cracked product stream in first cracked product line 46. The light olefin stream in light olefin line 106 comprises an oligomerization feed stream having C4 hydrocarbons including C4 olefins and perhaps having C5 hydrocarbons including C5 olefins.

An FCC recycle stream in recycle line 280 delivers an FCC recycle stream to the FCC zone 20. The FCC recycle stream is directed into a first FCC recycle line 202 with the control valve 202′ thereon opened. In an aspect, the FCC recycle stream may be directed into an optional second FCC recycle line 204 with the control valve 204′ thereon opened. The first FCC recycle line 202 delivers the first FCC recycle stream to the first FCC reactor 22 in an aspect to the riser 26 at an elevation above the first hydrocarbon feed in line 24. The second FCC recycle line 204 delivers the second FCC recycle stream to the second FCC reactor 70. Typically, both control valves 202′ and 204′ will not be opened at the same time, so the FCC recycle stream goes through only one of the first FCC recycle line 202 and the second FCC recycle line 204. However, feed through both is contemplated.

The second FCC recycle stream may be fed to the second FCC reactor 70 in the second FCC recycle line 204 via feed distributor 72. The second FCC reactor 70 may include a second riser 74. The second FCC recycle stream is contacted with catalyst delivered to the second riser 74 by a catalyst return pipe 76 to produce cracked upgraded products. The catalyst may be fluidized by inert gas such as steam from distributor 78. Generally, the second FCC reactor 70 may operate under conditions to convert the second FCC recycle stream to second cracked products such as ethylene and propylene. A second reactor vessel 80 is in downstream communication with the second riser 74 for receiving second cracked products and catalyst from the second riser 74. The mixture of gaseous, second cracked product hydrocarbons and catalyst continues upwardly through the second reactor riser 74 and is received in the second reactor vessel 80 in which the catalyst and gaseous, second cracked products are separated. A pair of disengaging arms may tangentially and horizontally discharge the mixture of gas and catalyst from a top of the second reactor riser 74 through one or more outlet ports 82 (only one is shown) into the second reactor vessel 80 that effects partial separation of gases from the catalyst. The catalyst can drop to a dense catalyst bed within the second reactor vessel 80. Cyclones 84 in the second reactor vessel 80 may further separate catalyst from second cracked products. Afterwards, a second cracked product stream can be removed from the second reactor 84 through an outlet in a second cracked product line 86 in downstream communication with the second reactor riser 74. The second cracked product stream in line 86 is fed to the FCC recovery zone 100, preferably separately from the first cracked products to undergo separation and recovery of ethylene and propylene. Separated catalyst may be recycled via a recycle catalyst pipe 76 from the second reactor vessel 80 regulated by a control valve back to the second reactor riser 74 to be contacted with the second FCC recycle stream.

In some embodiments, the second FCC reactor 70 can contain a mixture of the first and second FCC catalysts as described above for the first FCC reactor 22. In one preferred embodiment, the second FCC reactor 70 can contain less than about 20 wt-%, preferably less than about 5 wt-% of the first FCC catalyst and at least 20 wt-% of the second FCC catalyst. In another preferred embodiment, the second FCC reactor 70 can contain only the second FCC catalyst, preferably a ZSM-5 zeolite.

The second FCC reactor 70 is in downstream communication with the regenerator vessel 30 and receives regenerated catalyst therefrom in line 88. In an embodiment, the first FCC reactor 22 and the second FCC reactor 70 both share the same regenerator vessel 30. Line 90 carries spent catalyst from the second reactor vessel 80 to the lower chamber 54 of the regenerator vessel 30. The catalyst regenerator is in downstream communication with the second FCC reactor 70 via line 90.

The same catalyst composition may be used in both reactors 22, 70. However, if a higher proportion of the second FCC catalyst of small to medium pore zeolite is desired in the second FCC reactor 70 than the first FCC catalyst of large pore zeolite, replacement catalyst added to the second FCC reactor 70 may comprise a higher proportion of the second FCC catalyst. Because the second FCC catalyst does not lose activity as quickly as the first FCC catalyst, less of the second catalyst inventory must be forwarded to the catalyst regenerator 30 in line 90 from the second reactor vessel 80, but more catalyst inventory may be recycled to the riser 74 in return conduit 76 without regeneration to maintain a high level of the second FCC catalyst in the second reactor 70.

The second reactor riser 74 can operate in any suitable condition, such as a temperature of about 425° C. to about 705° C., preferably a temperature of about 550° C. to about 600° C., and a pressure of about 140 to about 400 kPa, preferably a pressure of about 170 to about 250 kPa. Typically, the residence time of the second reactor riser 74 can be less than about 3 seconds and preferably is than about 1 second. Exemplary risers and operating conditions are disclosed in, e.g., US 2008/0035527 A1 and U.S. Pat. No. 7,261,807 B2.

Before cracked products can be fed to the oligomerization zone 130, the light olefin stream in light olefin line 106 may require purification. Many impurities in the light olefin stream in light olefin line 106 can poison an oligomerization catalyst. Carbon dioxide and ammonia can attack acid sites on the catalyst. Sulfur containing compounds, oxygenates, and nitriles can harm oligomerization catalyst. Acetylene and diolefins can polymerize and produce gums on the catalyst or equipment. Consequently, the light olefin stream which comprises the oligomerization feed stream in light olefin line 106 may be purified in an optional purification zone 110.

The light olefin stream in light olefin line 106 may be introduced into an optional mercaptan extraction unit 112 to remove mercaptans to lower concentrations. In the mercaptan extraction unit 112, the light olefin feed may be prewashed in an optional prewash vessel containing aqueous alkali to convert any hydrogen sulfide to sulfide salt which is soluble in the aqueous alkaline stream. The light olefin stream, now depleted of any hydrogen sulfide, is contacted with a more concentrated aqueous alkali stream in an extractor vessel. Mercaptans in the light olefin stream react with the alkali to yield mercaptides. An extracted light olefin stream lean in mercaptans passes overhead from the extraction column and may be mixed with a solvent that removes COS in route to an optional COS solvent settler. COS is removed with the solvent from the bottom of the settler, while the overhead light olefin stream may be fed to an optional water wash vessel to remove remaining alkali and produce a sulfur depleted light olefin stream in line 114. The mercaptide rich alkali from the extractor vessel receives an injection of air and a catalyst such as phthalocyanine as it passes from the extractor vessel to an oxidation vessel for regeneration. Oxidizing the mercaptides to disulfides using a catalyst regenerates the alkaline solution. A disulfide separator receives the disulfide rich alkaline from the oxidation vessel. The disulfide separator vents excess air and decants disulfides from the alkaline solution before the regenerated alkaline is drained, washed with oil to remove remaining disulfides and returned to the extractor vessel. Further removal of disulfides from the regenerated alkaline stream is also contemplated. The disulfides are run through a sand filter and removed from the process. For more information on mercaptan extraction, reference may be made to U.S. Pat. No. 7,326,333 B2.

In order to prevent polymerization and gumming in the oligomerization reactor that can inhibit equipment and catalyst performance, it is desired to minimize diolefins and acetylenes in the light olefin feed in line 114. Diolefin conversion to monoolefin hydrocarbons may be accomplished by selectively hydrogenating the sulfur depleted stream with a conventional selective hydrogenation reactor 116. Hydrogen may be added to the purified light olefin stream in line 118.

The selective hydrogenation catalyst can comprise an alumina support material preferably having a total surface area greater than 150 m²/g, with most of the total pore volume of the catalyst provided by pores with average diameters of greater than 600 angstroms, and containing surface deposits of about 1.0 to 25.0 wt-% nickel and about 0.1 to 1.0 wt. % sulfur such as disclosed in U.S. Pat. No. 4,695,560. Spheres having a diameter between about 0.4 and 6.4 mm ( 1/64 and ¼ inch) can be made by oil dropping a gelled alumina sol. The alumina sol may be formed by digesting aluminum metal with an aqueous solution of approximately 12 wt-% hydrogen chloride to produce an aluminum chloride sol. The nickel component may be added to the catalyst during the sphere formation or by immersing calcined alumina spheres in an aqueous solution of a nickel compound followed by drying, calcining, purging and reducing. The nickel containing alumina spheres may then be sulfided. A palladium catalyst may also be used as the selective hydrogenation catalyst.

The selective hydrogenation process is normally performed at relatively mild hydrogenation conditions. These conditions will normally result in the hydrocarbons being present as liquid phase materials. The reactants will normally be maintained under the minimum pressure sufficient to maintain the reactants as liquid phase hydrocarbons which allow the hydrogen to dissolve into the light olefin feed. A broad range of suitable operating pressures therefore extends from about 276 (40 psig) to about 5516 kPa gauge (800 psig). A relatively moderate temperature between about 25° C. (77° F.) and about 350° C. (662° F.) should be employed. The liquid hourly space velocity of the reactants through the selective hydrogenation catalyst should be above 1.0 hr⁻¹. Preferably, it is between 5.0 and 35.0 hr⁻¹. The ratio of hydrogen to diolefinic hydrocarbons may be maintained between 0.75:1 and 1.8:1. The hydrogenation reactor is preferably a cylindrical fixed bed of catalyst through which the reactants move in a vertical direction.

A purified light olefin stream depleted of sulfur containing compounds, diolefins and acetylenes exits the selective hydrogenation reactor 116 in line 120. The optionally sulfur and diolefin depleted light olefin stream in line 120 may be introduced into an optional nitrile removal unit such as a water wash unit 122 to reduce the concentration of oxygenates and nitriles in the light olefin stream in line 120. Water is introduced to the water wash unit in line 124. An oxygenate and nitrile-rich aqueous stream in line 126 leaves the water wash unit 122 and may be further processed. A drier may follow the water wash unit 122.

A purified light olefin oligomerization feed stream perhaps depleted of sulfur containing compounds, diolefins and/or oxygenates and nitriles is provided in oligomerization feed stream line 128. The light olefin oligomerization feed stream in line 128 may be obtained from the cracked product stream in lines 46 and/or 86, so it may be in downstream communication with the FCC zone 20. The oligomerization feed stream need not be obtained from a cracked FCC product stream but may come from another source. The selective hydrogenation reactor 116 is in upstream communication with the oligomerization feed stream line 128. The oligomerization feed stream may comprise C4 hydrocarbons such as butenes, i.e., C4 olefins, and butanes. Butenes include normal butenes and isobutene. The oligomerization feed stream in oligomerization feed stream line 128 may comprise C5 hydrocarbons such as pentenes, i.e., C5 olefins, and pentanes. Pentenes include normal pentenes and isopentenes. Typically, the oligomerization feed stream will comprise about 20 to about 80 wt-% olefins and suitably about 40 to about 75 wt-% olefins. In an aspect, about 55 to about 75 wt-% of the olefins may be butenes and about 25 to about 45 wt-% of the olefins may be pentenes. Ten wt-%, suitably 20 wt-%, typically 25 wt-% and most typically 30 wt-% of the oligomerization feed may be C5 olefins.

The oligomerization feed line 128 feeds the oligomerization feed stream to an oligomerization zone 130 which may be in downstream communication with the FCC recovery zone 100.

In an aspect, an oligomerate return stream in oligomerate return line 231 to be described hereinafter may be mixed with the oligomerization feed stream in oligomerization feed conduit 128 in a first mixed oligomerization feed line 133. The oligomerization feed stream in line 133 may comprise about 10 to about 80 wt-% olefins and suitably about 25 to about 40 wt-% olefins if the oligomerate return stream from oligomerate return line 231 is mixed with the oligomerization feed stream. The oligomerization feed stream in line 133 may comprise about 10 to about 50 wt-% light olefins and suitably about 25 to about 40 wt-% light olefins if the oligomerate return stream from oligomerate return line 231 is mixed with the oligomerization feed stream. Accordingly, the oligomerization feed stream may comprise no more than about 38 wt-% butene and in another aspect, the oligomerization feed stream may comprise no more than about 23 wt-% pentene. The oligomerization feed stream to the oligomerization zone 130 in mixed oligomerization feed conduit 133 may comprise at least about 10 wt-% butene, and at least about 5 wt-% pentene. In a further aspect, the oligomerization feed stream may comprise no more than about 0.1 wt-% propylene. At least about 40 wt-% of the butene in the oligomerization feed stream may be normal butene. In an aspect, it may be that no more than about 70 wt-% of the oligomerization feed stream is normal butene. At least about 40 wt-% of the pentene in the oligomerization feed stream may be normal pentene. In an aspect, no more than about 70 wt-% of the oligomerization feed stream in the mixed oligomerization feed conduit 133 may be normal pentene.

The oligomerization reactor zone 130 comprises a first oligomerization reactor 138. The first oligomerization reactor may be preceded by an optional guard bed for removing catalyst poisons that is not shown. The first oligomerization reactor 138 contains the oligomerization catalyst. An oligomerization feed stream may be preheated before entering the first oligomerization reactor 138 in an oligomerization reactor zone 130. The first oligomerization reactor 138 may contain a first catalyst bed 142 of oligomerization catalyst. The first oligomerization reactor 138 may be an upflow reactor to provide a uniform feed front through the catalyst bed, but other flow arrangements are contemplated. In an aspect, the first oligomerization reactor 138 may contain an additional bed or beds 144 of oligomerization catalyst. C4 olefins in the oligomerization feed stream oligomerize over the oligomerization catalyst to provide an oligomerate comprising C4 olefin dimers and trimers. C5 olefins that may be present in the oligomerization feed stream oligomerize over the oligomerization catalyst to provide an oligomerate comprising C5 olefin dimers and trimers and co-oligomerize with C4 olefins to make C9 olefins. The oligomerization produces other oligomers with additional carbon numbers.

We have found that adding C5 olefins to the feed to the oligomerization reactor reduces oligomerization to heavier, distillate range material. This is counterintuitive since one may expect heavier C5 olefins to lead to the formation of more distillate range material. However, when C5 olefins dimerize with themselves or co-dimerize with C4 olefins, the C9 olefins and C10 olefins produced do not continue to oligomerize as quickly as C8 olefins produced from C4 olefin dimerization. Thus, the amount of net gasoline produced can be increased. In addition, the resulting C9 olefins and C10 olefins in the product have a very high octane value.

Oligomerization effluent from the first bed 142 may optionally be quenched with a liquid such as recycled oligomerate before entering the additional bed 144, and/or oligomerization effluent from the additional bed 144 of oligomerization catalyst may also be quenched with a liquid such as recycled oligomerate to avoid excessive temperature rise. The liquid oligomerate may also comprise oligomerized olefins that can react with the C4 olefins and C5 olefins in the feed and other oligomerized olefins if present to make diesel range olefins. Oligomerized product, also known as oligomerate, exits the first oligomerization reactor 138 in line 146.

In an aspect, the oligomerization reactor zone may include one or more additional oligomerization reactors 150. The oligomerization effluent may be heated and fed to the optional additional oligomerization reactor 150. It is contemplated that the first oligomerization reactor 138 and the additional oligomerization reactor 150 may be operated in a swing bed fashion to take one reactor offline for maintenance or catalyst regeneration or replacement while the other reactor stays online. In an aspect, the additional oligomerization reactor 150 may contain a first bed 152 of oligomerization catalyst. The additional oligomerization reactor 150 may also be an upflow reactor to provide a uniform feed front through the catalyst bed, but other flow arrangements are contemplated. In an aspect, the additional oligomerization reactor 150 may contain an additional bed or beds 154 of oligomerization catalyst. Remaining C4 olefins in the oligomerization feed stream oligomerize over the oligomerization catalyst to provide an oligomerate comprising C4 olefin dimers and trimers. Remaining C5 olefins, if present in the oligomerization feed stream, oligomerize over the oligomerization catalyst to provide an oligomerate comprising C5 olefin dimers and trimers and co-oligomerize with C4 olefins to make C9 olefins. Over 90 wt-% of the C4 olefins in the oligomerization feed stream can oligomerize in the oligomerization reactor zone 130. Over 90 wt-% of the C5 olefins in the oligomerization feed stream can oligomerize in the oligomerization reactor zone 130. If more than one oligomerization reactor is used, conversion is achieved over all of the oligomerization reactors 138, 150 in the oligomerization reactor zone 130.

Oligomerization effluent from the first bed 152 may be quenched with a liquid such as recycled oligomerate before entering the additional bed 154, and/or oligomerization effluent from the additional bed 154 of oligomerization catalyst may also be quenched with a liquid such as recycled oligomerate to avoid excessive temperature rise. The recycled oligomerate may also comprise oligomerized olefins that can react with the C4 olefins and C5 olefins in the feed and other oligomerized olefins to increase production of diesel range olefins.

An oligomerate conduit 156, in communication with the oligomerization reactor zone 130, withdraws an oligomerate stream from the oligomerization reactor zone 130. The oligomerate conduit 156 may be in downstream communication with the first oligomerization reactor 138 and the additional oligomerization reactor 150.

The oligomerization reactor zone 130 may contain an oligomerization catalyst. The oligomerization catalyst may comprise a zeolitic catalyst. The zeolite may comprise between 5 and 95 wt % of the catalyst. Suitable zeolites include zeolites having a structure from one of the following classes: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, MOR, OFF, MTW, TON, MTT, AFO, ATO, and AEL. In a preferred aspect, the oligomerization catalyst may comprise a zeolite with a framework having a ten-ring pore structure. Examples of suitable zeolites having a ten-ring pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO and FER. In a further preferred aspect, the oligomerization catalyst comprising a zeolite having a ten-ring pore structure may comprise a uni-dimensional pore structure. A uni-dimensional pore structure indicates zeolites containing non-intersecting pores that are substantially parallel to one of the axes of the crystal. The pores preferably extend through the zeolite crystal. Suitable examples of zeolites having a ten-ring uni-dimensional pore structure may include MTT. In a further aspect, the oligomerization catalyst comprises an MTT zeolite.

The oligomerization catalyst may be formed by combining the zeolite with a binder, and then forming the catalyst into pellets. The pellets may optionally be treated with a phosphoric reagent to create a zeolite having a phosphorous component between 0.5 and 15 wt % of the treated catalyst. The binder is used to confer hardness and strength on the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania and combinations of these metal oxides, and other refractory oxides, and clays such as montmorillonite, kaolin, palygorskite, smectite and attapulgite. A preferred binder is an aluminum-based binder, such as alumina, aluminum phosphate, silica-alumina and clays.

One of the components of the catalyst binder utilized in the present invention is alumina. The alumina source may be any of the various hydrous aluminum oxides or alumina gels such as alpha-alumina monohydrate of the boehmite or pseudo-boehmite structure, alpha-alumina trihydrate of the gibbsite structure, beta-alumina trihydrate of the bayerite structure, and the like. A suitable alumina is available from UOP LLC under the trademark Versal. A preferred alumina is available from Sasol North America Alumina Product Group under the trademark Catapal. This material is an extremely high purity alpha-alumina monohydrate (pseudo-boehmite) which after calcination at a high temperature has been shown to yield a high purity gamma-alumina.

A suitable oligomerization catalyst is prepared by mixing proportionate volumes of zeolite and alumina to achieve the desired zeolite-to-alumina ratio. In an embodiment, about 5 to about 80, typically about 10 to about 60, suitably about 15 to about 40 and preferably about 20 to about 30 wt-% MTT zeolite and the balance alumina powder will provide a suitably supported catalyst. A silica support is also contemplated.

Monoprotic acid such as nitric acid or formic acid may be added to the mixture in aqueous solution to peptize the alumina in the binder. Additional water may be added to the mixture to provide sufficient wetness to constitute a dough with sufficient consistency to be extruded or spray dried. Extrusion aids such as cellulose ether powders can also be added. A preferred extrusion aid is available from The Dow Chemical Company under the trademark Methocel.

The paste or dough may be prepared in the form of shaped particulates, with the preferred method being to extrude the dough through a die having openings therein of desired size and shape, after which the extruded matter is broken into extrudates of desired length and dried. A further step of calcination may be employed to give added strength to the extrudate. Generally, calcination is conducted in a stream of air at a temperature from about 260° C. (500° F.) to about 815° C. (1500° F.). The MTT catalyst is not selectivated to neutralize surface acid sites such as with an amine.

The extruded particles may have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most often have a symmetrical cross-sectional shape, preferably a spherical, cylindrical or polylobal shape. The cross-sectional diameter of the particles may be as small as 40 μm; however, it is usually about 0.635 mm (0.25 inch) to about 12.7 mm (0.5 inch), preferably about 0.79 mm ( 1/32 inch) to about 6.35 mm (0.25 inch), and most preferably about 0.06 mm ( 1/24 inch) to about 4.23 mm (⅙ inch).

The oligomerization reaction conditions in the oligomerization reactors 138, 150 in the oligomerization reactor zone 130 are set to keep the reactant fluids in the liquid phase. With liquid oligomerate recycle, lower pressures are necessary to maintain liquid phase. Operating pressures include between about 2.1 MPa (300 psia) and about 10.5 MPa (1520 psia), suitably at a pressure between about 2.1 MPa (300 psia) and about 6.9 MPa (1000 psia) and preferably at a pressure between about 2.8 MPa (400 psia) and about 4.1 MPa (600 psia). Lower pressures may be suitable if the reaction is kept in the liquid phase.

For the zeolite catalyst, the temperature in the oligomerization reactor zone 130 expressed in terms of a maximum bed temperature is in a range between about 150° C. and about 300° C., or about 160° C. and about 270° C., or about 180° C. and about 260° C., or about 200° C. and about 250° C. The space velocity should be between about 0.5 and about 5.0 hr⁻¹.

Across a single bed of oligomerization catalyst, the exothermic reaction will cause the temperature to rise. Consequently, the oligomerization reactor should be operated to allow the temperature at the outlet to be over about 25° C. greater than the temperature at the inlet. There will be a slight difference in the reactor temperature rise with an increase in temperature rise as the operation moves from maximum gasoline to maximum distillate. However, a unit designed for operation in the maximum distillate mode can easily be operated in the maximum gasoline mode.

The oligomerization reactor zone 130 with the oligomerization catalyst can be run in high conversion mode of greater than 95% conversion of feed olefins to produce a high quality distillate product and gasoline product. Normal butene conversion can exceed about 80%. Additionally, normal pentene conversion can exceed about 80%.

We have found that when C5 olefins are present in the oligomerization feed stream, they dimerize or co-dimerize with other olefins, but tend to mitigate further oligomerization over the zeolite with a 10-ring uni-dimensional pore structure. The highest mitigation of further oligomerization occurs when the C5 olefins comprise between 30 and 70 wt-% and preferably between 40 and 60 wt-% of the olefins in the oligomerization feed. Consequently, the oligomerate stream in oligomerate conduit 156 may comprise less than about 80 wt-% C9+ hydrocarbons when C5 olefins are present in the oligomerization feed at these proportions. Moreover, the oligomerate may comprise less than about 60 wt-% C12+ hydrocarbons when C5 olefins are present in the oligomerization feed at these proportions. Furthermore, the net gasoline yield may be at least about 40 wt-% when C5 olefins are present in the oligomerization feed.

If distillate is desired, however, the oligomerization zone with the oligomerization catalyst can be operated to oligomerize light olefins; i.e., C4 olefins, to distillate-range material by over 70 wt-% yield per pass through the oligomerization reactor zone 130. In an aspect, at least about 70 wt-% of the olefins in the oligomerization feed convert to C9+ product oligomers boiling above about 150° C. (302° F.) cut point in a single pass through the oligomerization zone. The C9+ oligomer from the oligomerization zone boiling above about 150° C. (392° F.) may have a cetane of at least 30 and preferably at least 40. The C12+ oligomer from the oligomerization zone boiling above about 200° C. (392° F.) may have a cetane of at least 30 and preferably at least 40.

The composition of the oligomerate in oligomerate line 156 may be an olefinic hydrocarbon composition having C8 olefins. The olefinic hydrocarbon composition may include gasoline. In an embodiment, the composition may be moderate in Type 2 disubstituted olefins and high in Type 4 trisubstituted olefins. In an aspect, the oligomerate composition may have a ratio of Type 2 disubstituted C8 olefins to Type 1 monosubstituted C8 olefins of greater than about 2. In a further aspect, a fraction of Type 2 disubstituted C8 olefins in the total C8 olefins in the oligomerate may be no less than about 7 and less than about 18 wt-%. In an even further aspect, a fraction of Type 4 trisubstituted C8 olefins in the total C8 olefins may be no less than about 40 wt-%. In a still further aspect, an average branch per C8 hydrocarbon molecule in the oligomerate may be less than 2. The oligomerate may have a cetane of greater than 30 and preferably greater than 40. The oligomerate may have a density of less than 0.83 kg/m³, preferably less than 0.81 kg/m³, less than 20 ppmw sulfur or less than 1 vol-% aromatics. The oligomerate may have a ratio of trimethylpentene to the total C8 olefins of no more than 50 and preferably no more than 40.

An oligomerization recovery zone 200 is in downstream communication with the oligomerization zone 130 and the oligomerate conduit 156. The oligomerate conduit 156 removes the oligomerate stream from the oligomerization zone 130.

The oligomerization recovery zone 200 may include a column 210 which separates the oligomerate stream into a first vaporous oligomerate overhead light stream comprising unreacted C4 and/or C5 olefins in a first overhead line 212 and a first liquid oligomerate bottom stream comprising an olefinic C5+ hydrocarbon stream in a first bottom line 214.

When maximum production of distillate is desired, the overhead pressure in the column 210 may be between about 400 and about 550 kPa (gauge) and the bottom temperature may be between about 250° C. and about 350° C.

When maximum production of gasoline is desired, the overhead pressure in the column 210 may be between about 400 and about 700 kPa (gauge) and the bottom temperature may be between about 250° C. and about 300° C.

The first vaporous oligomerate overhead light stream 212 comprising C4 and C5 olefins may be rejected from the process and subjected to further processing to recover useful components.

The first liquid oligomerate bottom stream comprising C5+ olefins and hydrocarbons in the first bottom line 214 can be apportioned into line 216 and line 218 with control thereon by valve 218′. C5+ olefinic hydrocarbon comprising effluent in line 216 is sent to a hydrogenation zone 220 where the olefins are hydrogenated with hydrogen 222 to paraffins. Suitable hydrogenation zones include, but are not limited to, any of the known hydrogenation technologies such as a standard hydrotreating zone or preferably it may be a hydrogenation zone specifically tailored for the hydrogenation of olefins, such as a complete saturation process. The hydrogenation zone is a catalytic zone and as such comprises a hydrogenation or hydrotreating catalyst. Hydrogenation or hydrotreating catalysts are any of those well known in the art such as nickel or nickel/molybdenum dispersed on a high surface area support. Catalysts which are commonly used for hydrotreating include those with molybdenum or tungsten as the active species in combination with other transition metal compounds such as nickel or cobalt. Usually these catalysts are dispersed on some type of support in order to maximize the effectiveness of the metals. Additionally, other additives such as phosphorus can be incorporated into the support either to stabilize the surface area of the support or to prevent formation of compounds such as cobalt or nickel aluminate. Other hydrogenation catalysts comprise a metal component selected from nickel and one or more noble metal catalytic elements dispersed on a high surface area support. Non-limiting examples of noble metals include Pt and/or Pd dispersed on gamma-alumina. Best results are obtained when substantially all of this metal component is present in the elemental state. This component may be present in the final catalyst composite in any amount which is catalytically effective, and generally will comprise about 0.01 to 2 mass-% of the final catalyst calculated on an elemental basis. Excellent results are obtained when the catalyst contains from about 0.05 to 1 mass-% of platinum. Hydrogenation conditions include a temperature of about 20° C. to about 400° C. and a pressure of about 0.3 MPa absolute (50 psia) to about 5.2 MPa absolute (750 psia) and a liquid hourly space velocity of from about 0.5 to about 50. A preferred embodiment is the use of a complete saturation process as the hydrogenation process. Suitable catalyst for this process will completely saturate mono- and polyolefinic hydrocarbons without significant cracking or polymerization activity. As such, the hydrogenated product stream can include less than about 0.1%, by weight, alkenes, and preferably less than about 0.01%, by weight, alkenes. Operating conditions include a temperature of about 20 to about 80° C. and a pressure of about 2 to about 3.5 MPa. An exemplary noble metal can be palladium.

If control valve 218′ is at least partially open, line 218 containing non-hydrotreated oligomerate product may be recycled to the oligomerization zone 130. The liquid oligomerate bottom product stream in the bottom product line 214 may have the same composition as described for the C8 olefins of the oligomerate in oligomerate line 156. The liquid oligomerate bottom product stream in the bottom product line 214 may have greater than 10 wt-% C10 isoolefins.

The operation of the hydrogenation zone will vary slightly from the maximum gasoline mode to the maximum distillate mode with less make-up hydrogen 222 required for the maximum distillate mode. A hydrogenation zone designed for maximum distillate mode can run in the maximum gasoline mode (and any mode between the two) because the only adjustment will be in the make-up hydrogen requirements.

The hydrotreated oligomerate 223 can be separated into a light ends stream 224 comprising unreacted hydrogen and any light paraffins in a stripper 225 and sent for further processing.

The hydrotreated oligomerate comprising C5+ paraffinic hydrocarbons exits the stripper 225 in line 226 which may be apportioned into lines 228 and line 230 with control from valve 230′ thereon. Line 228 is sent to the distillate separator column 240, while line 230 is recycled to the oligomerization zone 130.

It is desired to maintain liquid phase in the oligomerization reactors. Liquid phase may be maintained in the oligomerization zone 130 by incorporating into the feed recycle of the first liquid bottoms stream in line 218 from the column 210 and/or the hydrotreated oligomerate in line 230. Recycle line 231 is connected to lines 218 and 230. Valves 218′ and 230′ control the amount of flow through lines 218 and 230. The total amount of recycle in line 231 remains constant. The relative flow rates in line 218 and 230 will vary depending on whether maximum gasoline, maximum distillate, or something in between is desired. The recycle in line 231 can contain 0 to 100% of the unsaturated first liquid bottoms stream 218 (non-hydrotreated oligomerate product) and 0 to 100% of the hydrotreated oligomerate 230, depending on the desired product slate.

In an embodiment, the oligomerization zone 130 is in downstream communication with the first liquid bottoms stream 218 (non-hydrotreated oligomerate product) of the column 210. In a further aspect, the recycle hydrotreated oligomerate product line 230 and the oligomerate return line 231 are in downstream communication with the oligomerization zone 130. Consequently, the oligomerization zone 130 may be in upstream and downstream communication with the first bottom line 214, the recycle oligomerate product line 230 and the oligomerate return line 231. When maximum distillate product is desired, valve 230′ is closed, and the recycle in line 231 is entirely the first liquid bottoms stream in line 218 (non-hydrotreated oligomerate product).

When maximum gasoline product is desired, valve 218′ is closed, and the recycle in line 231 is entirely the hydrotreated oligomerate in line 230.

When a product slate between the two is desired, the recycle in line 231 will be a mixture of the first liquid bottoms stream in line 218 and the hydrotreated oligomerate in line 230, which is controlled by adjusting valves 218′ and 230′. The mixture can be at least about 95 wt % first liquid bottoms stream (non-hydrotreated oligomerate product), or at least about 90 wt %, or at least about 85 wt %, or at least about 80 wt %, or at least about 75 wt %, or at least about 70 wt %, or at least about 65 wt %, or at least about 60 wt %, or at least about 55 wt %, or at least about 50 wt %. The mixture can be at least about 95 wt % hydrotreated oligomerate, or at least about 90 wt %, or at least about 85 wt %, or at least about 80 wt %, or at least about 75 wt %, or at least about 70 wt %, or at least about 65 wt %, or at least about 60 wt %, or at least about 55 wt %, or at least about 50 wt %.

Increasing the amount of first liquid bottoms stream 218 (non-hydrotreated oligomerate product) increases the amount of distillate produced because the olefinic products are still active to oligomerize to higher olefins in the oligomerization zone. Increasing the amount of hydrotreated oligomerate 230 increases the amount of gasoline because the olefins have been hydrogenated and cannot further react; they serve to maintain the oligomerization zone 130 in the liquid phase.

The oligomerization recovery zone 200 may also include a distillate separator column 240 to which the distillate separator oligomerate feed stream comprising hydrotreated oligomerate may be fed in distillate feed line 228 for further separation.

The distillate separator column 240 separates the distillate separator oligomerate feed stream into an gasoline overhead stream in an overhead line 242 comprising C6, C7, C8, C9, C10 and/or C11 paraffinic hydrocarbons and a bottom distillate stream comprising C8+, C9+, C10+, C11+, or C12+ paraffinic hydrocarbons in a distillate bottom line 244. The distillate product stream in distillate bottom line 244 may be subjected to further processing to recover useful components or blended in the distillate pool. The distillate product stream may be in upstream communication with a distillate tank or a distillate blending line of a distillate pool. Additionally, LCO from LCO line 107 may also be blended with distillate in distillate bottom line 244 to form at least a portion of the overall refinery distillate pool.

When maximum production of distillate is desired, the overhead pressure in the distillate separator column 240 may be between about 10 and about 60 kPa (gauge) and the bottom temperature may be between about 225 and about 275° C.

When maximum production of gasoline is desired, the overhead pressure in the distillate separator column 240 may be between about 10 and about 60 kPa (gauge) and the bottom temperature may be between about 190 and about 250° C. The bottom temperature can be adjusted between about 175 and about 275° C. to adjust the bottom product between a C9+ cut and a C12+ cut based on the heaviness of the diesel cut desired by the refiner.

Surprisingly, altering the relative amounts of recycle flow back to the oligomerization reactors, 138 and 150 of the oligomerization reactor zone 130 with oligomerization catalyst significantly alters product selectivity and thus allows a flexible unit operation to produce varied amounts of gasoline and distillate. In an embodiment, the process may produce greater than 50 wt % distillate or greater than 60 wt % distillate or greater than 70 wt % distillate or greater than 80 wt % distillate or even greater than 85 wt % distillate. The first liquid oligomerate bottom stream, 214, may comprise greater than 50 wt % distillate or greater than 60 wt % distillate or greater than 70 wt % distillate or greater than 80 wt % distillate or even greater than 85 wt % distillate, the remainder being gasoline. The hydrotreated oligomerate, 226, may comprise greater than 50 wt % distillate or greater than 60 wt % distillate or greater than 70 wt % distillate or greater than 80 wt % distillate or even greater than 85 wt % distillate, the remainder being gasoline.

The first liquid oligomerate bottom stream may comprise C6-C11 olefins and preferably C7-C9 olefins and most preferably C8 olefins that can oligomerize with C4-05 olefins in the oligomerization feed stream in the oligomerization zone 130 to diesel range material comprising C10-C16 diesel product. C4 olefins continue to oligomerize with C4 olefins and C5 olefins if present in the feed.

The oligomerization catalyst, and particularly, the uni-dimensional, 10-ring pore structured zeolite, converts a significant fraction of the gasoline-range olefins, such as C8 olefins, to distillate material by oligomerizing them with feed olefins, such as C4 and/or C5 olefins. Additionally, the presence of the gasoline-range olefins also encourages oligomerization of the feed olefins with each other over the zeolite catalyst. Surprisingly, the isobutene conversion is lower than normal butene conversion at high overall butene conversion such as over 90% C4 olefin conversion. When gasoline is recycled from the first liquid bottoms stream in line 218 to the oligomerization reactor zone 130 for oligomerization over uni-dimensional, 10-ring pore structured zeolite, oligomerate from the oligomerization zone in oligomerate line 156 may comprise greater than 30 wt-% C9+ olefins. Under these circumstances, oligomerate from the oligomerization reactor zone in oligomerate line 156 may comprise greater than 50 wt-% or even greater than 60 wt-% C9+ olefins.

When production of less distillate product is required, recycling hydrotreated oligomerate comprising C5+ paraffinic hydrocarbons may be preferred. With less olefinic gasoline comprising the feed in oligomerate return line 231, oligomerization of feed light olefins with gasoline range olefins may be discouraged and higher gasoline selectivities achieved.

In an aspect, the gasoline overhead stream in gasoline overhead line 242 may be recovered as product in product gasoline line 248 in downstream communication with the recovery zone 200. A control valve 248′ may be used to completely shut off flow through gasoline product line 248 or allow partial or full flow therethrough. The gasoline product stream may be subjected to further processing to recover useful components or blended in the gasoline pool. The gasoline product line 248 may be in upstream communication with a gasoline tank 252 or a gasoline blending line of a gasoline pool. In this aspect, the overhead line 242 of the distillate separator column may be in upstream communication with the gasoline tank 252 or the gasoline blending line.

In an embodiment, the bottom stream in a bottom line 218 or 230 may be recycled to the FCC zone 20 in FCC recycle line 280 via a recycle oligomer line 233 in downstream communication with the oligomerization recovery zone 200 to be cracked to propylene product in the FCC zone. The recycle oligomer stream may comprise C5+, C6+, C9+, C10+, C11+ or C12+ olefins that can crack to propylene. In a preferred aspect of this embodiment, the recycle oligomer stream comprises the liquid bottoms stream 218. A control valve 233′ may be used to completely shut off flow through recycle oligomer line 233 or allow partial or full flow therethrough. In this embodiment, the FCC zone 20 is in downstream communication with the separator column 210 and particularly the liquid bottoms line 214.

The recycle oligomer stream subjected to FCC is converted best over a blend of medium or smaller pore zeolite blended with a large pore zeolite such as Y zeolite as explained previously with respect to the FCC zone 20. Additionally, oligomerate produced over the oligomerization catalyst in the oligomerization reactor zone 140 provides an excellent feed to the FCC zone that can be cracked to yield greater quantities of propylene.

Preferred embodiments of this invention are described herein, including the best mode known to the inventors for carrying out the invention. It should be understood that the illustrated embodiments are exemplary only, and should not be taken as limiting the scope of the invention.

Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.

In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated. Pressures are given at the vessel outlet and particularly at the vapor outlet in vessels with multiple outlets. Control valves should be opened or closed as consistent with the intent of the disclosure.

From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.

The invention will now be further illustrated by the following non-limiting examples.

Examples

The feeds in Table 1 were contacted with Catalyst A to determine how the composition of the recycle affected the oligomer product. Contacting was performed in a ½″ steel reactor using 28 g of catalyst at 900 psig, max bed temperature 236° C. and 1.5 WHSV.

TABLE 1 Feed A Feed B Feed C Component Fraction, wt-% Fraction, wt-% Fraction, wt-% propane 3.71 3.71 2.7 propylene 10.68 10.68 10.8 isobutane 13.5 13.5 13.5 n-butane 0.06 0.06 0.3 isobutylene 8.21 8.21 7.8 1-butene 3.55 3.55 3.3 2-butene 9.90 9.90 8.1 (cis and trans) 2-methylbutane 0.05 0.05 1.4 n-hexane 49.92 31.52 0 other hexanes 0.45 0.29 0 C6+ olefin recycle 0 18.53 52.1

Catalyst A is an MTT catalyst purchased from Zeolyst having a product code ZD2K019E and extruded with alumina to be 25 wt-% zeolite. Extrusion and calcination were performed according to procedures described in US20140135554. The MTT catalyst was not selectivated to neutralize acid sites such as with an amine.

The results are shown in Table 2.

TABLE 2 Feed A Feed B Feed C representing representing representing 100% Stream 37/63 Mix 100% Stream 230 Maximum Streams 218 and 218 Maximum Description Gasoline 230 Mix Mode Distillate Conversion n-C4 = 86 86 92 conversion, wt % i-C4 = 95 96 96 conversion, wt % C3 = 94 96 97 conversion, wt % Selectivity C5-C8 Sel, % 48 17 14 C9+ Sel, % 52 83 86 C12+ Sel, % 45 76 78

Selectivities are defined as the amount of compound appearing in the GC boiling between the n-alkane 1 carbon number less than the starting Cn listed and the n-alkane of the same carbon number as the Cn listed with the amount of feed in the area subtracted, the number divided by the conversion. Thus, for C5-C8 selectivity, these are compounds which boil between n-butane and n-octane with any hexanes comprising the feed subtracted divided by the total conversion of olefins in the feed.

The surprising results above indicate that by recycling different streams to the oligomerization zone to form a portion of the oligomerization zone feed, varied yields of gasoline and distillate can be made over Catalyst A, MTT. This is confirmed by the greater net yield of gasoline and the lower selectivity to C12+ fraction for Feed A than for Feed B and for Feed B than for Feed C. Thus recycling greater proportions of paraffinic streams to the oligomerization zone cause higher yields of gasoline range product. Also, by recycling olefinic streams back to the oligomerization zone feed a greater yield of distillate range material can be made. This is confirmed by the greater net yield of distillate for Feed C than for Feed B on a single pass basis and for Feed B than for Feed A. In this way, the yield of gasoline versus distillate can be modified to any proportion desired by the individual operator. Gasoline yield was classified by product meeting the Engler T90 requirement and distillate yield was classified by product boiling over 150° C. (300° F.).

While at least one exemplary embodiment has been presented in the foregoing detailed description of the invention, it should be appreciated that a vast number of variations exist. It should also be appreciated that the exemplary embodiment or exemplary embodiments are only examples, and are not intended to limit the scope, applicability, or configuration of the invention in any way. Rather, the foregoing detailed description will provide those skilled in the art with a convenient road map for implementing an exemplary embodiment of the invention. It being understood that various changes may be made in the function and arrangement of elements described in an exemplary embodiment without departing from the scope of the invention as set forth in the appended claims. 

What is claimed:
 1. A process for controlling an amount of gasoline and distillate obtained from an oligomerate product comprising: introducing an oligomerization feed stream comprising at least one of C4 or C5 olefins into an oligomerization zone under oligomerization conditions in the presence of an oligomerization catalyst to oligomerize olefins in the oligomerization feed stream in liquid phase to produce a mixture of an oligomerate product comprising C5+ hydrocarbons and unreacted C4 or C5 feed olefins; and separating the mixture into a light stream comprising the unreacted C4 or C5 feed olefins and a stream of the oligomerate product in a first separation zone; optionally dividing the stream of oligomerate product into a first portion and a second portion, the first portion being a non-hydrotreated oligomerate product; hydrogenating at least the second portion of the stream of the oligomerate product to produce a stream of hydrotreated oligomerate product; introducing a recycle stream into the oligomerization zone, the recycle stream comprising at least one of the optional first portion of non-hydrotreated oligomerate product and a recycle portion of the stream of hydrotreated oligomerate product; and controlling a flow rate of the optional first portion of non-hydrotreated oligomerate product and a flow rate of the recycle portion of the stream of hydrotreated oligomerate product in the recycle stream to control the amount of gasoline and distillate product produced.
 2. The process of claim 1 wherein the optional first portion of non-hydrotreated oligomerate product is present and wherein controlling the flow rate of the first optional portion of non-hydrotreated oligomerate product and the flow rate of the portion of the stream of hydrotreated oligomerate product in the recycle stream comprises increasing the flow rate of the portion of the hydrotreated oligomerate product in the recycle stream to increase the amount of the gasoline product produced.
 3. The process of claim 1 wherein the optional first portion of non-hydrotreated oligomerate product is present and wherein controlling the flow rate of the first optional portion of non-hydrotreated oligomerate product and the flow rate of the portion of the hydrotreated oligomerate product in the recycle stream comprises increasing the flow rate of the first portion of non-hydrotreated oligomerate product in the recycle stream to increase the amount of the distillate product produced.
 4. The process of claim 1 wherein the recycle stream comprises at least about 50 wt % of hydrotreated oligomerate product and less than about 50 wt % of non-hydrotreated oligomerate product.
 5. The process of claim 1 wherein the recycle stream comprises at least about 90 wt % of hydrotreated oligomerate product and less than about 10 wt % of non-hydrotreated oligomerate product.
 6. The process of claim 1 wherein the recycle stream comprises at least about 90 wt % of non-hydrotreated oligomerate product and less than 10 wt % of hydrotreated oligomerate product.
 7. The process of claim 1 wherein the oligomerization conditions include an operating temperature in a range of about 160° C. to about 260° C. and a pressure in a range of about 2.1 MPa (g) to about 10.5 MPa (g).
 8. The process of claim 1 wherein the oligomerization catalyst comprises a zeolite with a framework having a ten-ring pore structure.
 9. The process of claim 8 wherein the oligomerization catalyst comprises a zeolite having a uni-dimensional, ten-ring pore structure.
 10. The process of claim 9 wherein the oligomerization catalyst comprises an MTT zeolite.
 11. The process of claim 1 wherein the oligomerization feed stream is obtained by: cracking a hydrocarbon feed over an FCC catalyst in an FCC zone to produce spent FCC catalyst and a cracked product stream; and obtaining the oligomerization feed stream from the cracked product stream.
 12. A process for controlling an amount of gasoline and distillate obtained from an oligomerization product comprising: cracking a hydrocarbon feed over an FCC catalyst in an FCC zone to produce spent FCC catalyst and a cracked product stream; obtaining an oligomerization feed stream comprising at least one of C4 or C5 olefins from the cracked product stream; introducing the oligomerization feed stream into an oligomerization zone under oligomerization conditions in the presence of an oligomerization catalyst to oligomerize olefins in the oligomerization feed stream in liquid phase to produce a mixture of an oligomerate product comprising C5+ hydrocarbons and unreacted C4 or C5 feed olefins; separating the mixture into a light stream comprising the unreacted C4 or C5 olefins and a stream of oligomerate product in a first separation zone; optionally dividing the stream of oligomerate product into a first portion and a second portion, the first portion being a non-hydrotreated oligomerate product; hydrogenating at least the second portion of the stream of oligomerate product to produce a stream of hydrotreated oligomerate product; dividing the stream of hydrotreated oligomerate into a product portion and a recycle portion; separating a product portion of the stream of hydrotreated oligomerate product into a stream of gasoline and a stream of distillate in a second separation zone; introducing a recycle stream into the oligomerization zone, the recycle stream comprising at least one of the of the optional first portion of non-hydrotreated oligomerate product and the recycle portion of the hydrotreated oligomerate product; and controlling a flow rate of the optional first portion of the non-hydrotreated stream of oligomerate product and a flow rate of the recycle portion of the hydrotreated oligomerate product in the recycle stream to control the amount of gasoline and distillate produced; wherein controlling the flow rate of the optional first portion of the non-hydrotreated oligomerate product and the flow rate of the recycle portion of the hydrotreated oligomerate product in the recycle stream comprises increasing the flow rate of the recycle portion of the hydrotreated oligomerate product in the recycle stream to increase the amount of gasoline product produced or increasing the flow rate of the optional first portion of the non-hydrotreated oligomerate product in the recycle stream to increase the amount of distillate product produced.
 13. The process of claim 12 wherein the recycle stream comprises at least about 50 wt % of hydrotreated oligomerate product and less than about 50 wt % of non-hydrotreated oligomerate product.
 14. The process of claim 12 wherein the recycle stream comprises at least about 90 wt % of hydrotreated oligomerate product and less than about 10 wt % of non-hydrotreated oligomerate product.
 15. The process of claim 12 wherein the recycle stream comprises at least about 90 wt % of the non-hydrotreated stream of oligomerate product and less than about 10 wt % of hydrotreated oligomerate product.
 16. The process of claim 12 wherein the oligomerization conditions include an operating temperature in a range of about 160° C. to about 260° C.
 17. The process of claim 12 wherein the oligomerization catalyst comprises a zeolite with a framework having a ten-ring pore structure.
 18. The process of claim 12 wherein the oligomerization catalyst comprises a zeolite having a uni-dimensional, ten-ring pore structure.
 19. The process of claim 12 wherein the oligomerization catalyst comprises an MTT zeolite.
 20. An apparatus for controlling the production of diesel and gasoline comprising: an oligomerization reaction zone with an inlet and an outlet; a first separation zone with an inlet, an overhead outlet, and a bottoms outlet; the inlet of the separation zone being in fluid communication with the outlet of the oligomerization zone; a hydrogenation zone having an inlet and an outlet, the inlet of the hydrogenation zone being in fluid communication with the bottoms outlet of the separation zone; a second separation zone with an inlet, an overhead outlet, and a bottoms outlet, the inlet of the second separation zone being in fluid communication with the outlet of the hydrogenation zone; the bottoms outlet of the first separation zone being in fluid communication with the inlet of the oligomerization zone and the outlet of the hydrogenation zone being in fluid communication with the inlet of the oligomerization zone; and a controller for controlling a flow from the bottoms outlet of the first separation zone to the inlet of the oligomerization zone and a flow from the outlet of the hydrogenation zone to the inlet of the oligomerization zone. 